Conversion of co-fed methane and hydrocarbon feedstocks into higher value hydrocarbons

ABSTRACT

In one aspect, the inventive process comprises a process for pyrolyzing a hydrocarbon feedstock containing nonvolatiles in a regenerative pyrolysis reactor system. The process comprises: (a) heating the nonvolatile-containing hydrocarbon feedstock upstream of a regenerative pyrolysis reactor system to a temperature sufficient to form a vapor phase that is essentially free of nonvolatiles and a liquid phase containing the nonvolatiles; (b) separating said vapor phase from said liquid phase; (c) feeding the separated vapor phase and methane to the pyrolysis reactor system; and (d) converting the methane and separated vapor phase in said pyrolysis reactor system to form a pyrolysis product. In another aspect, the invention includes a separation process that feeds multiple pyrolysis reactors.

RELATIONSHIP TO OTHER APPLICATIONS

This application claims benefit of U.S. provisional application Ser. No.60/933,011, filed Jun. 4, 2007, and claims benefit of U.S. provisionalapplication Ser. No. 60/933,044, filed Jun. 4, 2007, the entirety ofwhich are incorporated herein by reference.

FIELD OF THE INVENTION

This invention pertains to conversion of hydrocarbons using regenerativepyrolysis reactors. The invention relates to a process for crackinghydrocarbons present in hydrocarbon feedstocks containing nonvolatiles,in a regenerative pyrolysis reactor. The nonvolatiles are removed fromthe feedstocks before the hydrocarbons undergo thermal pyrolysis. Moreparticularly the invention relates to a process and apparatus forimproving the quality of nonvolatile-containing feedstocks to aregenerative pyrolysis reactor system or a plurality of such pyrolysisreactor systems.

BACKGROUND OF THE INVENTION

Conventional steam crackers are known as an effective tool for crackinghigh-quality feedstocks that contain a large fraction of volatilehydrocarbons, such as ethane, gas oil, and naphtha. Similarly,regenerative pyrolysis reactors are also known and conventionally usedfor converting or cracking and to execute cyclic, high temperaturechemistry such as those reactions that may be performed at temperatureshigher than can suitably be performed in conventional steam crackers.Regenerative reactor cycles typically are either symmetric (samechemistry or reaction in both directions) or asymmetric (chemistry orreaction changes with step in cycle). Symmetric cycles are typicallyused for relatively mild exothermic chemistry, examples beingregenerative thermal oxidation (“RTO”) and autothermal reforming(“ATR”). Asymmetric cycles are typically used to execute endothermicchemistry, and the desired endothermic chemistry is paired with adifferent chemistry that is exothermic (typically combustion) to provideheat of reaction for the endothermic reaction. Examples of asymmetriccycles are Wulff cracking, Pressure Swing Reforming, and otherregenerative pyrolysis reactor processes. Regenerative pyrolysisreactors are generally known in the art as being capable of convertingor cracking hydrocarbons. However, they have not achieved commercial orwidespread use for hydrocarbon conversion, due at least in part to thefact that they have not scaled well to an economical size. This failureis in large part due to the inability of the equipment to adequatelycontrol and contend with the very high temperatures and the way thatfuel and oxidant are combined during the regeneration or heating stageof the process. The high temperatures are difficult to position andcontain for extended periods of time and lead to premature equipmentfailure. A solution was proposed in U.S. patent application Ser. No.11/643,541 filed in the USPTO, on Dec. 21, 2006, entitled “MethaneConversion to Higher Hydrocarbons,” related primarily to methanefeedstocks for pyrolysis systems, utilizing an inventive deferredcombustion process.

As with steam crackers, regenerative pyrolysis reactors also are wellsuited for volatized or volatizable feedstocks that are substantiallyfree of nonvolatile components, such as metals and other residual ornonvolatizable components, which would otherwise lay down and build upin the reactor as ash. Pyrolysis reactors typically operate at highertemperatures than steam crackers. Nonvolatiles may be defined broadly tomean any resid, metal, mineral, ash-forming, asphaltenic, tar, coke, orother component or contaminant within the feedstock that will notvaporize below a selected boiling point or temperature and which, duringor after pyrolysis, may leave an undesirable residue or ash within thereactor system. The nonvolatile components of most concern are thosethat deposit as ash within the reactor system and cannot be easilyremoved by regeneration. Many hydrocarbon coke components may be merelyburned out of the bed at the high temperature typically used inpyrolysis reactor systems and thus tend to be of less concern than someother residual components. Some nonvolatile feed components, such asmetals and/or minerals, may leave an ash component behind that even atthe high regeneration temperatures is difficult to remove from areactor.

Typically, regenerative reactors include a reactor bed or zone,typically comprising some type of refractory material, where thereaction takes place within the reactor system. Conventionalregenerative reactors typically deliver a stream of fuel, oxidant, or asupplemental amount of one of these reactants, directly to a locationsomewhere within the flow path of the reactor bed. The deliveredreactants then are caused to exothermically react therein and heat thereactor media or bed. Thereafter, the reacted reactants are exhaustedand a pyrolysis feedstock, such as a hydrocarbon feed stream, preferablyvaporized, is introduced into the heated region of the reactor media orbed, and exposed to the heated media to cause heating and pyrolysis ofthe reactor feedstock into a pyrolyzed reactor feed. The pyrolyzedreactor feed is then removed from the reaction area of the reactor andquenched or cooled, such as in a quench region of the reactor system, tohalt the pyrolysis reaction and yield a pyrolysis product.

However, as with steam cracking, economics may favor using lower costfeedstocks such as, by way of non-limiting examples, crude oil, heavydistillate cuts, contaminated naphthas and condensates, and atmosphericresids, as feedstocks for regenerative pyrolysis reactors.Unfortunately, these economically favored feedstocks typically containundesirable amounts of nonvolatile components and have heretofore beenunacceptable as regenerative reactor feedstocks. The nonvolatiles leadto fouling of the reactor through deposition of materials such as ash,metals, and/or coke. Regenerative pyrolysis reactors do not have theflexibility to process such otherwise economically crack favorablefeedstocks because, although coke can typically be burned off, depositsor buildup of ash and metals within the reactor cannot easily be burnedor removed. The critical concentration of nonvolatiles within aparticular feedstock may vary depending upon variables such as theintended process, feedstock conditions or type, reactor design,operating parameters, etc. Generally, nonvolatile concentrations (e.g.,ash, metals, resids, etc.) in excess of 2 ppmw (ppm by weight) of thefeed stream to the reactor will cause significant fouling in a pyrolysisreactor. Some economically desirable lower cost feeds may contain up to10 percent by weight of nonvolatiles, while still other feeds maycontain well in excess of 10 weight percent of nonvolatiles. Sincenonvolatiles do not vaporize, but decompose to form ash, metals, tar,and/or coke when heated above about 600° F. (315° C.) (in an oxidizingenvironment), the nonvolatiles present in disadvantaged feedstocks laydown or build up as a foulant in the reaction section of pyrolysisreactors, which increases pressure drop through the reactor and leads toplugging and decreased efficiency. Generally, only low levels ofnonvolatiles (e.g., <2 ppmw and preferably <1 ppmw) or more specificallylow levels of ash (measured by ASTM D482-03 or ISO 6245:2001) can betolerated in the reactor feeds. Nonvolatiles are generally determined inaccordance with ASTM D6560.

Various techniques have been employed for treating petroleum hydrocarbonfeeds for the removal of nonvolatiles contained therein to render costadvantaged feeds suitable for conventional steam cracker feeds. Theseprocesses tend to improve the quality of hydrocarbon feeds containingnonvolatiles for conventional steam cracking. However, in most instancesthe processes suffer from operating condition limitations, spacelimitations for retrofits, high capital costs, and high operating costs,due to the processing steps used, high capital expense of the requiredequipment, and/or unsatisfactory reduction limitations in the amount ofnonvolatiles removed from the feeds. For example, it may be quite costlyto equip each of the several steam cracking furnaces in a steam crackingcomplex with all of the equipment necessary to process the low costfeedstocks to provide an acceptable, nonvolatile-free feed into thecracking section of each steam cracker. Similar and even exaggeratedproblems exist for a regenerative pyrolysis reactor complex, due totheir feed quality requirements and increased temperature severity.

The present invention provides a revolutionary process for improving thequality of nonvolatile-containing hydrocarbon feedstocks to render suchfeed suitable for use as a feedstream to a regenerative pyrolysisreactor system. The invention provides a commercially useful and costeffective technique for removing the ash-forming nonvolatiles from thefeedstock before the feedstock undergoes pyrolysis in a regenerativepyrolysis reactor.

SUMMARY OF THE INVENTION

The present invention relates to pyrolysis of hydrocarbons and in oneembodiment includes a process for reducing ash formation due topyrolyzing a hydrocarbon feedstock containing nonvolatiles, in aregenerative pyrolysis reactor system. In one aspect, the inventiveprocess comprises a process for reducing ash formation from pyrolyzing ahydrocarbon feedstock containing nonvolatiles in a regenerativepyrolysis reactor system, the process comprising: (a) heating thenonvolatile-containing hydrocarbon feedstock upstream of a regenerativepyrolysis reactor system (preferably a reverse flow type of regenerativereactor system) to a temperature sufficient to form a vapor phase(preferably essentially free of nonvolatiles) and a liquid phasecontaining nonvolatiles; (b) separating said vapor phase from saidliquid phase; (c) transferring or feeding at least a portion of theseparated vapor phase and methane to the pyrolysis reactor system; and(d) converting (e.g., cracking) at least a portion of the separatedvapor phase in the regenerative pyrolysis reactor system to form apyrolysis product. The methane may be added to the hydrocarbon feedstockand/or the separated vapor phase stream at substantially any point(s) inthe process, such as upstream of the separation process, within theseparation process, in the separated vapor phase transfer line, into areactor of the reactor system, and/or combinations thereof. The methanemay also be converted to pyrolysis product within the reactor systemand/or may complement conversion of the separated vapor phase. Inanother embodiment, the pyrolysis reactor system comprises at least twopyrolysis reactor systems and the separated vapor is cracked in at leasttwo of the at least two pyrolysis reactor systems. The separated vaporis combined with methane. The methane may be added or co-fed as asubstantially neat methane stream or associated with or included withinanother hydrocarbon feed stream containing methane, such as a mixedhydrocarbon stream, typically comprising at least 10 weight percentmethane, based upon the total stream weight. The separated vapor streamand the methane-containing stream may be combined upstream of theseparation unit, within the separation unit, upstream of the reactor,within the reactor system, and/or combinations thereof. Other diluentsmay also be added to the process to affect the desired chemistry. Thecombined methane and separated vapor-phase feed is converted in aregenerative reactor system, preferably in at least two or more reactorsystems by feeding the vapor in parallel flow (e.g., substantiallysimultaneously) to the at least two reactor systems, to provide forcommercially useful quantities of pyrolysis product.

In another embodiment, the invention includes the step of feeding adiluent or stripping agent, such as hydrogen, to the pyrolysis reactorsystem in conjunction with the separated vapor phase and methane feed,for cracking the mixed vapor phase and methane in the presence of thediluent or stripping agent, within the regenerative pyrolysis reactorsystem. Hydrogen may typically be a preferred diluent or strippingagent. Steam may be used as a diluent or stripping agent for somealternative processes. However, steam may not be preferred in manyprocesses, in that, unlike steam cracking, at pyrolysis reactortemperatures, such as above 1200° C., steam can react with hydrocarbonto form carbon monoxide and hydrogen. However, depending upon the amountof methane present, the need for hydrogen diluent may be reduced oreliminated. The reaction of cracking the methane to acetylene will yieldfree hydrogen which can serve as a stripping agent or diluent.Therefore, the need for additional hydrogen may be reduced as the ratioof methane to separated vapor phase is increased.

In yet another aspect, the invention comprises an inventive process forthe manufacture of a hydrocarbon pyrolysis product, such as olefins,aromatics, and/or acetylene, from the methane and vaporized hydrocarbonfeed using a reverse-flow type regenerative pyrolysis reactor system,wherein the reactor system includes (i) a first reactor comprising afirst end and a second end, and (ii) a second reactor comprising primaryend and a secondary end, the first and second reactors oriented in aseries flow relationship with respect to each other such that thesecondary end of the second reactor is proximate the second end of thefirst reactor. In one aspect, the inventive process comprises the stepsof: (a) heating a nonvolatile-containing hydrocarbon feedstock upstreamof the regenerative pyrolysis reactor system to a temperature sufficientto form a vapor phase that is essentially free of nonvolatiles and aliquid phase containing the nonvolatiles; (b) separating the vapor phasefrom the liquid phase; (c) supplying a first reactant through a firstchannel in the first reactor and supplying at least a second reactantthrough a second channel in the first reactor, such that the first andsecond reactants are supplied to the first reactor from the first end ofthe first reactor; (d) combining the first and second reactants at thesecond end of the first reactor and reacting the combined reactants toproduce a heated reaction product; (e) passing the heated reactionproduct through the second reactor to transfer heat from the reactionproduct to the second reactor to produce a heated second reactor; (f)transferring at least a portion of the separated vapor phase from step(b) and methane to the pyrolysis reactor system, whereby the separatedvapor phase and methane commingled (e.g., combined or mixed) with eachother within the reactor system; (g) feeding the separated vapor phaseand methane through the heated second reactor to convert the separatedvapor phase and methane into a pyrolysis product; (h) quenching thepyrolysis product; and (i) recovering the quenched pyrolysis productfrom the reactor system.

BRIEF DESCRIPTION OF THE DRAWINGS

FIGS. 1( a) and 1(b) are a simplified, diagrammatic illustration of thetwo steps in a regenerating reverse flow pyrolysis reactor systemaccording to the present invention.

FIG. 2 is another diagrammatic illustration of an exemplary regenerativebed reactor system that defers combustion, controls the location of theexothermic reaction, and adequately quenches the recuperation reactorbed.

FIG. 3 illustrates an axial view of an exemplary gas distributor.

FIG. 4 illustrates a cross sectional view of an exemplary gas/vapormixer and channels for controlled combustion. FIG. 4 a is a cutout viewof a portion of FIG. 4.

FIG. 5 is a simplified process flow diagram illustrating an embodimentof the invention.

DETAILED DESCRIPTION

The terms “convert” and “converting” are defined broadly herein toinclude any molecular decomposition, cracking, breaking apart,conversion, and/or reformation of organic molecules in the hydrocarbonfeed, by means of at least pyrolysis heat, and may optionally includesupplementation by one or more of catalysis, hydrogenation, diluents,and/or stripping agents.

As used herein, the expression “essentially free of nonvolatiles” meansthat concentration of nonvolatiles in the vapor phase is reduced to anextremely low level. Those skilled in the art know that it is difficultto obtain a complete separation of nonvolatiles from a hydrocarbonfeedstock such as crude oil. As a result, the vapor phase may contain atrace amount of nonvolatiles. Therefore, in the context of the presentinvention, while it is the objective that the vapor phase contains nononvolatiles, it is recognized that the vapor phase may contain anacceptable trace amount of nonvolatiles, e.g., typically an amount of 2ppmw or less, but still be considered essentially free of nonvolatiles.The separated vapor phase preferably contains less than 1 ppmw ofnonvolatiles. More preferably, the vapor phase contains less than 0.5ppmw of nonvolatiles. Variables such as the pyrolysis conditions andreactor design will dictate an appropriate threshold cutoff fornonvolatile carryover in the vapor phase, for a specific application.

FIG. 5, illustrates a simplified schematic flow diagram of anon-limiting embodiment of the invention, including feeding ortransferring methane (2) and a hydrocarbon feedstock that containsnonvolatile components therein via inlet line (1) to a heat unit/zone(3). Preferred methane to hydrocarbon feed molar ratios may range fromabout 1 to about 5. That is, the amount of methane added to thehydrocarbon feedstock or to the separated vapor phase preferably resultsin a methane to separated vapor phase molar ratio of from about 1:1 toabout 5:1, upon entering the heated, pyrolysis reaction section of thereactor system (e.g., entering the region of the reactor where thepyrolysis reaction chemistry or conversion occurs). Substantially, anyhydrocarbon feedstock containing a mixture of both volatiles andnonvolatiles can advantageously be utilized in the process. Examples ofsuch feedstock include one or more of steam cracked gas oil andresidues, gas oils, heating oil, jet fuel, diesel, kerosene, gasoline,coker naphtha, steam cracked naphtha, catalytically cracked naphtha,hydrocrackate, reformate, raffinate reformate, Fischer-Tropsch liquids,Fischer-Tropsch gases, natural gasoline, distillate, naphtha, crude oil,atmospheric pipestill bottoms, vacuum pipestill streams includingbottoms, virgin naphtha, wide boiling range naphthas, heavy non-virginhydrocarbon streams from refineries, vacuum gas oil, heavy gas oil,naphtha contaminated with crude, atmospheric resid, heavy residuum,C₄'s/residue admixture, condensate, contaminated condensate, naphtharesidue admixture and mixtures thereof. The hydrocarbon feedstock mayhave a nominal end boiling point of at least 400° F. (200° C.), (e.g.,greater than or equal to 400° F., such as in excess of 1200° F. and evenin excess of 1500° F.) and will commonly have a nominal end boilingpoint of at least 500° F. (260° C.). Some preferred hydrocarbonfeedstocks include crude oil, atmospheric resids, contaminatedcondensate, and gas oil distillates, tars, fuel oils and cycle oils. Themethane and vaporized hydrocarbon feed may include substantially anyother hydrocarbon co-feed material that undergoes the endothermicreforming, such as to acetylene, including natural gas mixtures, otherpetroleum alkanes, petroleum distillates, kerosene, jet fuel, fuel oil,heating oil, diesel fuel and gas oil, gasoline, and alcohols. Apreferred co-feed may be a hydrocarbon component that may function as ahydrogen donor diluent, such as tetralin, and dihydroanthracene,hydropyrene, or hydrotreated steam cracked tar oils. Preferably, thefeed will be in a vapor or gaseous state at the temperature and pressureof introduction into the reactor system.

Hydrocarbon streams that have been processed through a refinery, e.g.,naphtha, gas oils, etc., may be suitable reactor feeds. The inventiveprocess will cleanse the feeds if the same become contaminated withash-forming components in their transport to the pyrolysis facility.However, use of typical light refiner streams may limit theattractiveness of the process due to the relatively high cost of thefeeds, as compared to other heavier feeds. Heavier, more aromatic feedsare typically lower cost, per unit weight, but may yield lower acetyleneand ethylene yields and higher carbon or tar yields. Due to the higharomatic content of the heavier feeds, the feeds have lower hydrogencontent and during pyrolysis, the hydrogen deficit feeds may form tar,coke, or soot. Conversion of the methane co-feed during reforming willprovide additional hydrogen to facilitate better conversion of theseparated vapor phase from the heavier feeds.

The amount of nonvolatiles present in the hydrocarbon feedstock willvary depending upon the feedstock source and quality. For example,contaminates, full range vacuum gas oils, and petroleum crude oils oftencontain relatively high levels of nonvolatile molecules, for example, upto 20 percent by weight of nonvolatiles. Other feedstocks may containeven higher concentrations of nonvolatiles. A typical hydrocarbonfeedstock used in the process of the present invention may containnonvolatiles in an amount of from about 5 to about 40 weight percentbased upon the weight of the total hydrocarbon feed. Feeding a hydrogendiluent or stripping agent should help offset this deficit to facilitateproduction of preferred products, such as ethylene, acetylene, hydrogen,and methane.

In heat unit (3), the hydrocarbon feedstock is heated to a temperaturethat is sufficient to form a vapor phase and a liquid phase. The heatingof the hydrocarbon feedstock is not limited to any particular technique.For example, the heating can be conducted by means such as, but notlimited to, a heat exchanger, steam injection, submerged heat coil, or afired heater. In some embodiments, the heat unit may be a separate unit,such as illustrated by element (3) in FIG. 5, and in other embodimentsthe heat unit may be integrated with or internal to separation unit (7).The temperature to which the hydrocarbon feedstock is heated will varydepending upon composition of the hydrocarbon feedstock, methane and/orhydrogen concentration, and the desired cut-off point for distinguishingthe vaporized fraction and the liquid fraction. Commonly, thenonvolatile-containing hydrocarbon feedstock comprises a liquid phaseand the feedstock is heated to a temperature at which at least 50percent of the liquid phase hydrocarbon feedstock is converted to avapor phase, preferably greater than 90 weight percent, and morepreferably greater than 98 weight percent of the feedstock is vaporized.Exemplary separation temperatures may range from 400° F. to 1200° F.(200° C. to 650° C.). Preferably, the hydrocarbon feedstock is heated toa temperature from 450° F. to 1000° F. (230° C. to 540° C.), and morepreferably from 500° F. to 950° F. (260° C. to 510° C.). Since thenonvolatiles contained in the hydrocarbon feedstock are essentiallynonvolatile, they remain within the nonvolatized liquid phase. Thefraction of nonvolatiles in each of the vapor/liquid phases is afunction of both the hydrocarbon partial pressure and the temperature towhich the hydrocarbon feedstock is heated. Desirably, about 50 to about98 percent by weight of the heated feedstock will be in the vapor phase.Still more preferably, at least 90 weight percent of the feedstock willbe volatized into the vapor phase. Vaporizing substantially all of thefeedstock may become more difficult with heavier feedstocks. Foridentification purposes, the vaporized or volatized fraction of theseparated feed stream may be referred to herein as the separated vaporphase, even if such fraction is wholly or fully condensed, partiallycooled or condensed, stored, and/or later revaporized, prior to feedinginto the pyrolysis furnace. Preferably the separated vapor phase is fedto the pyrolysis furnace in a vapor/gas phase.

Referring still to FIG. 5, for embodiments having an external orseparate heat and separation units, the heated feedstock is transferredvia line (5) to a separation unit (7), where the vapor phase isseparated from the liquid phase. For integrated or internal heat units,the vapor phase is separated from the liquid phase in a vapor-liquidseparation unit. Examples of equipment suitable for separating the vaporphase from the liquid phase include knock-out drum (e.g., substantiallyany vapor-liquid separator), a flash drum, distillation column/unit,flash drum having a heating means within the drum, a knock-out drumhaving heating means within the known-out drum, and combinationsthereof. Exemplary heating means may include direct fired heaters,steam, convection heating, heat exchangers, radiant heating,electric-resistance heating, or other heat source. In many embodimentsit may be important to affect the feed stream separation step so thatthe vapor phase is essentially free of nonvolatiles, e.g., having lessthan 1 weight percent of nonvolatiles carried into the separated vaporphase, based upon the total weight of the separated vapor phase,determined substantially at or near the exit of the separation vessel.Otherwise, the nonvolatiles entrained in the vapor phase will be carriedinto the pyrolysis reactor and may cause coking and/or ash problems.

Heat unit (3) and separation unit (7) are located upstream with respectto the pyrolysis reactor system. Upstream merely means that thehydrocarbon feed is first separated into vapor and liquid phases, andthen the vapor phase is transferred to the pyrolysis reactors. There mayalso be intermediate steps or processes, such as, for example,introducing hydrogen into the vapor phase and/or hydrogenation of thevapor phase before cracking. Although the heat unit and separation unitare depicted in FIG. 5 as separate units, they can be combined into asingle unit (“heat/separation unit”). Examples of suitableheat/separation units include distillation towers, fractionators, andvisbreakers, as well as knock-out drums and flash drums having a meanswithin the drum for heating the hydrocarbon feedstock. Examples ofsuitable techniques for heating of the hydrocarbon feedstock containedwithin the heat/separation unit include injecting hydrogen into thehydrocarbon feedstock present in the heat/separation unit, heating in ahydrogenation unit/process, and heaters immersed into the liquidhydrocarbon feedstock present in the heat/separation unit. Additionallyand preferably, fired heaters may be used to heat the hydrocarbonfeedstock. Heating of the nonvolatile-containing hydrocarbon feedstockmay be carried out such as by fired heater, heat exchanger (eitherinternal or external, including but not limited to conventional heatexchangers, submerged internal coils or elements, convection or radiantheating, induction heating, and/or heat from the reaction system), steaminjection, and/or combinations thereof. Although the heat unit (3) andseparation unit (7) are each shown in FIG. 5 as respective single andseparate units, each of these units can alternatively comprise aplurality of units, e.g., a separation unit can include more than oneknock-out drums, separators, and/or flash drums. As discussed below, theheat unit (3) and separation unit (7) may also be combined or integratedinto substantially a common unit.

For some process embodiments, it may be preferred to maintain adetermined constant ratio of vapor to liquid within the separation unit(7) or, as the case may be, the heat/separation unit, but such ratio isdifficult to measure and control. However, the temperature of the heatedfeedstock before separation can be used as an indirect parameter tomeasure, control, and maintain an approximately constant vapor to liquidratio in the unit. Ideally, the higher the feedstock temperature, thehigher percentage of hydrocarbons that will be vaporized and becomeavailable as part of the vapor phase for cracking. However, when thefeedstock temperature is too high, nonvolatiles such as coke precursorscould be present in the vapor phase and carried over to the convectionreactor tubes, eventually coking and/or ashing the tubes. The hydrogendiluent and/or methane co-feed, however, will help suppress cokeprecursor formation or at least make it palatable, since the additionalfree hydrogen produced in the reactor will facilitate burning off thecoke in the reactor. Ashing in the reactor, however, should still beavoided. A primary objective of the feed separation step is to removeashing precursors. Conversely, if the temperature of the heatedfeedstock is too low, this can result in a low ratio of vapor to liquidwith more volatile hydrocarbons remaining in the liquid phase and not beavailable for cracking.

Adding the methane co-feed and/or hydrogen diluent to the separatorpermits raising the temperature of the separation step and vessel, ascompared to such separation step in the absence of hydrogen and/ormethane. Thereby, feeding the methane or other co-feed and optionallyhydrogen diluent or another hydrogen donor diluent (e.g., naphthalene),into the separator may enable volatizing a higher percentage of thehydrocarbon feed without formation of unmanageable ash/coke precursors,as compared to the absence of hydrogen diluent with the feed. However,the methane, steam, hydrogen, or other additives may also be introducedinto hydrocarbon stream at substantially any location in the processthat is upstream of the heated reaction area of the reactor. Forexample, such additives may be introduced upstream of the separationstep, during the separation step, after separation in the separatedvapor phase, into the reactor system and/or combinations thereof,depending upon the specific process chemistry used, operatingconditions, and feedstock properties.

The maximum separation temperature of the heated feedstock may alsodepend upon the composition of the hydrocarbon feedstock. If thefeedstock contains higher amounts of lighter hydrocarbons, theseparation temperature of the feedstock may be lower while vaporizing anacceptable percentage of the feed. If the feedstock contains a higheramount of less-volatile or higher boiling point hydrocarbons, thetemperature of the feedstock may be heated to a higher value forseparation, but may also need the hydrogen diluent. For example, withrespect to vacuum gas oil feeds, the temperature of the heatedfeedstream may typically be maintained in the range of from 400° F.(200° C.) to 1200° F. (650° C.).

In addition to temperature, it is usually also desirable to maintain asubstantially constant hydrocarbon partial pressure to maintain asubstantially constant ratio of vapor to liquid in the separationvessel. Typically, the hydrocarbon partial pressures for the heatedfeedstream are dependent upon the amount of hydrogen or other strippingagent present in or mixed with the feed. In one aspect of the inventiveprocess, it may be preferred to combine the methane with the hydrocarbonfeed stream either upstream of the separation step or directly into theseparation step. Similarly, hydrogen may also be added to thehydrocarbon stream upstream of or in the separation step. The methaneadded into the stream may assist improved vaporization and separation ofthe hydrocarbon feed in the heated separation step. Additional methaneand hydrogen may also be added to the separated vapor stream and/ordirectly into the pyrolysis reactor.

The amount of vapor phase produced in the separation step can varywidely, depending upon the application and feedstock input rate. Forexample, in some applications the vapor phase flow rate may be a vaporflow rate that has only a partial pyrolysis reactor load, while in otherapplications, the vapor flow rate may simultaneously load a plurality(two or more) of pyrolysis reactors. Still further, in some applicationsthe vapor phase flow rate may exceed the reactor(s) load capacity for aparticular installation, whereby the excess vapor cut may be condensedand stored for subsequent use in either steam cracking and/or aspyrolysis reactor feed, or sent to other applications or uses. Forexample, the condensed material can be stored for at least a day, week,or even longer, such as in tanks or other storage vessels, or sent asfeed to other processes. The determination of total pyrolysis reactorload capacity is determinable by persons skilled in the processing art.For example, total load capacity may be calculated from the heatrequirements, flow capacity, reaction requirements, etc. Pyrolysiscapacity is sometimes limited by the heat output capabilities of thereactor and efficiency with which that heat is utilized and movedthrough the reactor system. In some embodiments, the inventive processincludes using multiple pyrolysis reactors, such as at least twopyrolysis reactor systems, including at least a first pyrolysis reactorsystem, and the amount of separated vapor phase is in excess of thereactor capacity of the first pyrolysis reactor system. Thereby,additional reactors may be used to handle the total capacity of thetransferred vapor phase feed for pyrolysis. For example, a singleheater-separator system may feed two or more reactor systems, such as abank of reactor systems.

Referring still to FIG. 5, the nonvolatile-containing liquid phase maybe withdrawn or removed from separation zone (7) as a bottoms stream,such as via line (9). This material can be sold as fuel oil or furtherprocessed, e.g., subjected to fluidized catalytic cracking (FCC),coking, or POX to produce higher value products, etc. The liquid phasemay also contain resins in addition to nonvolatiles. Resins differ fromthe nonvolatiles primarily in having lower molecular weight, lesspolynuclear aromatics, more solubility in aliphatic hydrocarbons, andlower in metal content.

The separated vapor phase may be withdrawn from separation unit (7) asan overhead stream via line (11) and passed to one or a plurality (twoor more) of pyrolysis reactors, such as illustrated FIG. 5, depictingtwo reactors as pyrolysis reactor systems (17) and (19). The separated,vaporized hydrocarbons may include various concentrations of associatedgases, such as ethane and other alkanes. The vaporized fraction may alsoinclude impurities, such as H₂S and/or nitrogen, and may be sweetenedbefore feeding to the reactor system. Methane, including amethane-containing feed, (2) may be mixed (e.g., commingled, introduced,fed into, or otherwise combined) into the separated vapor phase line atsubstantially any point from, and including, the separator (7) to thepyrolysis unit(s) (17) (19). A convenient place for introduction of themethane-containing feed may be a transfer line (11), such as illustratedby methane feed line (2). If the methane source is a hydrocarbon streamthat comprises methane, the stream should comprise at least 10 weightpercent methane, preferably at least 30 weight percent methane, and morepreferably at least 50 weight percent methane. Other alkanes, such asethane, may also be present. Although two pyrolysis reactors areillustrated, three or more pyrolysis reactors also may be used in someapplications. Alternatively, the vapor phase essentially free ofnonvolatiles can be removed via line (21), cooled to a liquid in coolingunit (23), and then transferred via line (25) to storage unit (27).Although the cooling unit and storage unit are each depicted in FIG. 5as separate units, in other applications they may comprise a common(e.g., substantially integrated or combined) heat-separator unit. Acommon heater-separator unit may include, for example, one or more of adistillation column, a flash drum having a heating means within thedrum, a knock-out drum having a heating means within the knock-out drum,and combinations thereof. Some installations may also include aplurality of common units to serve one or more reactor systems.

Also, each of the cooling and unit and/or the storage unit may compriseone or more of such respective unit, e.g., storage unit can comprise aplurality of tanks. The liquid (or a portion thereof) can be transferredfrom storage unit (27) via line (29) to line (11) and then passed ortransferred in substantially parallel flow, such as via lines (13) and(15), to pyrolysis reactors (17) and (19). The cracked reaction productmay then be transferred to product-separation processes, such as viaoutlet lines (49) and (51).

As illustrated in exemplary FIG. 5, the vaporized phase or cut from theseparation in the separating unit (7) (either with or without interimstorage (27)) and the methane (2), may be transferred feed line (11) toone or more pyrolysis reactor systems, such as illustrated by reactorsystems (17) and (19), such as via lines (13) and (15). The methane (2)and separated vapor phase feed through lines (13) and/or (15) isintroduced into the respective reactor system(s) and heated to atemperature sufficient for conversion or cracking of the methane andvapor stream to a mix of higher value hydrocarbons, such as acetylenes.Alternatively, the co-fed methane may be introduced directly into thesecond reactor of each of the reactor systems (17) and (19), along withthe separated vapor phase.

According to a preferred process, the methane and separated vapor feedare exposed to the previously heated hot spot or reaction zone withinthe reactor system for a determined appropriate residence time(typically less than 1.0 second, commonly less than 0.5 seconds, andoften less than 0.1 seconds, while a preferred range of 1-100 ms ispreferred) and then quenched to stop the reaction to provide the desiredselectivity to a preferred hydrocarbon product mix or pyrolysis productwithin the cracked product stream. Longer reaction times tend to favorthe formation of coke. In many preferred applications, the reaction willbe allowed to proceed for sufficient time to crack the vapor phasehydrocarbons into smaller components, such as methyl groups (e.g., CH₄,CH₃, and CH₂) and hydride radicals. At least a portion of the methane isconverted to an acetylene pyrolysis product in the reactor system. Atleast a portion of the introduced or intermediately produced methane ormethyl radicals are converted to acetylene pyrolysis product in thereactor system. Aromatic molecules may similarly be converted toacetylenes or diacetylenes radicals pyrolysis product. A methane co-feedmay further help form hydride radicals, which may help suppress coke andhelp the reaction proceed to formation of acetylene. A typical preferredprocess may also include a relatively high selectivity (≧50 weightpercent) for acetylene within the final cracked product stream mix.Other exemplary products that may result from a preferred process mayinclude hydrogen and methane, along with some other components such asresidual coke. Some components of the vaporized feed stream may beconverted within the reactor system, directly to acetylene. For example,in a high severity regenerative reactor, the high temperature will startbreaking carbons or methyl radicals off of aliphatic or nonaromaticchains, while the aromatics within the feed may be reformed directly toacetylenes or diacetylenes. Preferably, sufficient quenching occurswithin the reactor system such that separate additional quench steps(e.g., heat exchangers, etc.) are not required to stop the conversionfrom running beyond the production of a high selectivity to acetylenepyrolysis product. The cracked product mix may include gaseoushydrocarbons of great variety, e.g., from methane to coke, and mayinclude saturated, monounsaturated, polyunsaturated, and aromatics. Insome aspects, the pyrolysis product produced is a dilute acetylenestream (primarily acetylene, with some hydrogen and unreacted methane)that can be easily hydrogenated to an olefin, such as ethylene, in thevapor phase or liquid phase. The acetylene hydrogenation reactor couldbe, for example, a standard fixed bed process using knownhydroprocessing catalyst.

In another exemplary process, a vaporized stream from the separationprocess may comprise a mix of hydrocarbons, such as aliphatic,naphthenic, and aromatic compounds. Such vapor stream may be condensedand stored for later feeding to a regenerative reactor system or fed tothe reactor system without substantially going through an intermediatecondensation step. The regenerative reactor may be heated according to aregeneration process whereby exothermically reacting components, such asfuel and oxidant are introduced and reacted in the reactor to heat thereactor media, with the resulting reaction product removed from thereactor. Then the vaporized feed and methane may be introduced into orpassed through the heated zone within the reactor.

Typical conditions may include a residence time from 0.001 to 1.0seconds and may typically include, for example, a pressure from about 5to 50 psia (34 to 345 kPa). In some embodiments, the reactor conditionsmay be at a vacuum pressure, such as less than 15 psia (103 kPa).Cracked pyrolysis product may be removed from the reactor system, suchas via lines 49 and/or 51 and transferred to other processes forrecovery of the various component products of the cracked product. Thereactor system may also include additional feed lines (not shown) suchas fuel and oxidant feed, stripping agent feed, exhaust lines, etc.

The regenerative pyrolysis reactor system according to this invention isgenerally a higher temperature hydrocarbon pyrolysis reactor system thantypical steam cracking type hydrocarbon systems that are conventionallyused in commercial steam cracking operations. For example, commercialnaphtha steam cracking operations typically operate at furnace radiantcoil outlet temperatures of less than about 815° C. (1500° F.). However,the terms “regenerative pyrolysis reactor systems” as pertaining to thesubject invention refers to cyclical (regenerating) thermal hydrocarbonpyrolysis systems that heat the hydrocarbon stream to be converted(e.g., the separated vapor phase) to temperatures of at least 1200° C.(2192° F.), preferably in excess of 1400° C. (2552° F.), more preferablyin excess of 1500° C. (2732° F.), or for some applications, even morepreferably in excess of 1700° C. (3092° F.). In some reactions, it mayeven be still more preferable to heat the feeds for very short timeduration, such as less than 0.1 seconds, to a temperature in excess of2000° C. (3632° F.). Pyrolysis reactions that benefit from reaction orconversion of the co-fed methane in addition to the hydrocarbon vapor,typically require reactor temperatures in excess of 1400° C. (2552° C.)for the methane to react or convert. An exemplary preferred process maypyrolyze the feed stream within the reactor, such as at temperatures offrom about 1500 to about 1900° C., and more preferably from about 1600to about 1700° C. Exemplary residency times preferably may be short,such as less than 0.1 seconds and preferably less than about 5milliseconds. In some aspects, the conversion or cracking of the methaneand separated vapor phase may be performed in the presence of hydrogen,hydride, other hydrocarbons, and/or other diluents or stripping agents.The conversion of the vapor fraction into higher value hydrocarbons suchas acetylene typically requires a high reformation temperature, which inthe past has been a significant barrier to commercialization andefficiency.

At least part of the invention of the present inventors is therecognition that the requisite high temperature may be achieved bycreating a high-temperature heat bubble in the middle of a packed bedsystem. This heat bubble may be created via a two-step process whereinheat is (1) added to the reactor bed via delayed, in-situ combustion,and then (2) removed from the bed via in-situ endothermic reforming. Akey benefit of the invention is the ability to consistently manage andconfine the high temperature bubble (e.g., >1600° C.) in a reactorregion(s) that can tolerate such conditions long term. The inventiveprocess provides for a substantially continuously operating,large-scale, cyclic, regenerative reactor system that is useful andoperable on a commercial scale, thereby overcoming the limitations ofthe prior art.

A regenerative reactor system or process may be described generally ashydrocarbon pyrolysis in a regenerative reactor or more specifically theconversion of a volatized hydrocarbon stream to acetylene or otherpyrolysis produce via thermal pyrolysis of the hydrocarbons in aregenerative reactor system. One exemplary regenerative pyrolysisreactor system includes first and second reactors and comprises areverse flow type of regenerative pyrolysis reactor system, such asillustrated in FIGS. 1( a) and 1(b). In one preferred arrangement, thefirst and second reactors may be oriented in a series flow relationshipwith each other, with respect to a common flow path, and more preferablyalong a common center axis. The common axis may be horizontal, vertical,or otherwise. A regenerative pyrolysis reactor is a cyclical reactorwhereby in a first part of the cycle materials may flow and react for aperiod of time in one direction through the reactor, such as to generateand transfer heat to the reactor media, and then in a second part of thecycle the same and/or other materials may be fed through the reactor toreact in response to the heat and thereby produce a pyrolysis product.In a reverse flow type of regenerative pyrolysis reactor, during thesecond portion of the cycle the materials flow in an opposite directionas compared to the direction of material flow in the first portion ofthe cycle. The regenerative pyrolysis reactor system contains a reactionzone that includes the heated or hot area of the reactor where themajority of the high temperature reaction chemistry takes place, and aquenching zone that serves to absorb heat from the reacted product andquench and halt the reaction process or chemistry by cooling thereaction product. At least a portion of the separated vapor feed (andpreferably at least a portion of the co-fed methane) that is transferredto or fed into the reactor system is, generally, (i) cracked in thereaction zone to form the pyrolysis product, and (ii) that crackedreaction product is timely quenched in the quenching zone to stop thereaction at the desired pyrolysis product step to thereby yield thepyrolysis product. If the reaction is not timely quenched, the reactionmay continue breaking the molecules into either coke, their elementalcomponents, or less desirable product components.

The present invention includes a process wherein first and secondin-situ combustion reactants are both separately, but preferablysubstantially simultaneously, passed through a quenching reactor bed(e.g., a first reactor bed), via substantially independent flow paths(channels), to obtain the quenching (cooling) benefits of the totalcombined weight of the first and second reactants. (Although only firstand second reactants are discussed, the regeneration reaction may alsoinclude additional reactants and reactant flow channels.) Both reactantsare also concurrently heated by the hot quench bed, before they reach adesignated location within the reactor system and react with each otherin an exothermic reaction zone (e.g., a combustion zone). This deferredcombustion of the first and second reactants permits positioninginitiation of the exothermic regeneration reaction, in-situ, at thedesired location within the reactor system.

The reactants are permitted to combine or mix in the reaction zone tocombust therein, in-situ, and create a high temperature zone or heatbubble (e.g., 1500-1700° C.) inside of the reactor system. Preferablythe combining is enhanced by a reactant mixer that mixes the reactantsto facilitate substantially complete combustion/reaction at the desiredlocation, with the mixer preferably located between the first and secondreactors. The combustion process takes place over a long enough durationthat the flow of first and second reactants through the first reactoralso serves to displace a substantial portion, (as desired) of the heatproduced by the reaction (e.g., the heat bubble), into and at leastpartially through the second reactor, but preferably not all of the waythrough the second reactor to avoid waste of heat and overheating thesecond reactor. The flue gas may be exhausted through the secondreactor, but preferably most of the heat is retained within the secondreactor. The amount of heat displaced into the second reactor during theregeneration step is also limited or determined by the desired exposuretime or space velocity that the volatized hydrocarbon feed gas will haveto the reforming, high temperature second reactor media to convert thevolatized hydrocarbon and other hydrocarbons to acetylene.

After regeneration or heating the second reactor media, in thenext/reverse step or cycle, the methane and volatized hydrocarbon cutfrom the previously discussed separation step are fed or flowed throughthe second reactor, preferably from the direction opposite the directionof flow during the heating step. The methane and volatized hydrocarbonscontact the hot second reactor and mixer media, in the heat bubbleregion, to transfer the heat to the volatized hydrocarbon for reactionenergy. In addition to not wasting heat, substantially overheating thereformer/second reactor bed may adversely lead to a prolonged reactionthat cracks the hydrocarbons past the acetylene-generation point,breaking it down into its elemental components. Thus, the total amountof heat added to the bed during the regeneration step should not exceedthe sum of the heats that are required (a) to sustain the reformingreaction for the endothermic conversion of the supplied hydrocarbon toacetylene for a suitable period of time, as determined by many factors,such as reactor size, dimensions, vapor flow rates, temperatures used,desired contact time, cycle duration, etc, and (b) for heat losses fromthe system both as conduction losses through reactor walls as well asconvective losses with the exiting products. The total amount of heatstored in the reactor system though is generally much more heat thanwould be minimally needed for conversion on any single cycle. However,it is desirable to avoid having the temperature bubble so large that theresidence time at temperature becomes too long. As is commonly done forreactor systems, normal experimentation and refining adjustments andmeasurements can be made to the reactor system to obtain the optimum setof reactor conditions.

In preferred embodiments, the reactor system may be described ascomprising two zones/reactors: (1) a heat recuperating (first)zone/reactor, and (2) a reforming (second) zone/reactor. As a catalystis preferably not required to facilitate reforming the hydrocarbon vaporto acetylene, so in most preferred embodiments, no catalyst is presentin the reactor beds. However, there may be some applications thatbenefit from the presence of a catalyst to achieve a certain range ofreforming performance and such embodiments are within the scope of theinvention.

The basic two-step asymmetric cycle of a regenerative bed reactor systemis depicted in FIGS. 1 a and 1 b in terms of a reactor system having twozones/reactors; a first or recuperator/quenching zone (7) and a secondor reaction/reforming zone (1). Both the reaction zone (1) and therecuperator zone (7) contain regenerative beds. Regenerative beds, asused herein, comprise materials that are effective in storing andtransferring heat. The term regenerative reactor means a regenerativemedia that may also be used for carrying out a chemical reaction. Theregenerative beds may comprise bedding or packing material such as glassor ceramic beads or spheres, metal beads or spheres, ceramic (includingzirconia) or metal honeycomb materials, ceramic tubes, extrudedmonoliths, and the like, provided they are competent to maintainintegrity, functionality, and withstand long term exposure totemperatures in excess of 1200° C. (2192° F.), preferably in excess of1500° C. (2732° F.), more preferably in excess of 1700° C. (3092° F.),and even more preferably in excess of 2000° C. (3632° F.) for operatingmargin.

As shown in FIG. 1 a, at the beginning of the “reaction” step of thecycle, a secondary end (5) of the reaction zone (1) (a.k.a. herein asthe reformer or second reactor) is at an elevated temperature ascompared to the primary end (3) of the reaction bed (1), and at least aportion (including the first end (9)) of the recuperator or quench zone(7), is at a lower temperature than the reaction zone (1) to provide aquenching effect for the synthesis gas reaction product. A hydrocarboncontaining reactant feed, and preferably also a diluent or strippingagent, such as hydrogen or steam, is introduced via a conduit(s) (15),into a primary end (3) of the reforming or reaction zone (1). Thereby,in one preferred embodiment, the term pyrolysis includes hydropyrolysis.

The feed stream from inlet(s) (15) absorbs heat from the reformer bed(1) and endothermically reacts to produce the desired acetylene product.As this step proceeds, a shift in the temperature profile (2), asindicated by the arrow, is created based on the heat transfer propertiesof the system. When the bed is designed with adequate heat transfercapability, this profile has a relatively sharp temperature gradient,which gradient will move across the reaction zone (1) as the stepproceeds. The sharper the temperature gradient profile, the better thereaction may be controlled.

The reaction gas exits the reaction zone (1) through a secondary end (5)at an elevated temperature and passes through the recuperator reactor(7), entering through a second end (11), and exiting at a first end (9)as a synthesized gas comprising acetylene, some unconverted methyls, andhydrogen. The recuperator (7) is initially at a lower temperature thanthe reaction zone (1). As the synthesized reaction gas passes throughthe recuperator zone (7), the gas is quenched or cooled to a temperatureapproaching the temperature of the recuperator zone substantially at thefirst end (9), which in some embodiments is preferably approximately thesame temperature as the regeneration feed introduced via conduit (19)into the recuperator (7) during the second step of the cycle. As thereaction gas is cooled in the recuperator zone (7), a temperaturegradient (4) is created in the zone's regenerative bed(s) and movesacross the recuperator zone (7) during this step. The quenching heatsthe recuperator (7), which must be cooled again in the second step tolater provide another quenching service and to prevent the size andlocation of the heat bubble from growing progressively through thequench reactor (7). After quenching, the reaction gas exits therecuperator at (9) via conduit (17) and is processed for separation andrecovery of the various components.

The second step of the cycle, referred to as the regeneration step, thenbegins with reintroduction of the first and second regenerationreactants via conduit(s) (19). The first and second reactants passseparately through hot recuperator (7) toward the second end (11) of therecuperator (7), where they are combined for exothermic reaction orcombustion in or near a central region (13) of the reactor system.

The regeneration step is illustrated in FIG. 1 b. Regeneration entailstransferring recovered sensible heat from the recuperator zone (7) tothe reaction zone (1) to thermally regenerate the reaction beds (1) forthe subsequent reaction cycle. Regeneration gas/reactants entersrecuperator zone (7) such as via conduit(s) (19), and flows through therecuperator zone (7) and into the reaction zone (1). In doing so, thetemperature gradients (6) and (8) may move across the beds asillustrated by the arrows on the exemplary graphs in FIG. 1( b), similarto but in opposite directions to the graphs of the temperature gradientsdeveloped during the reaction cycle in FIG. 1( a). Fuel and oxidantreactants may combust at a region proximate to the interface (13) of therecuperator zone (7) and the reaction zone (1). The heat recovered fromthe recuperator zone together with the heat of combustion is transferredto the reaction zone, thermally regenerating the regenerative reactionbeds (1) disposed therein.

In a preferred embodiment of the present invention, a first reactant,such as a hydrocarbon fuel, is directed down certain channels (eachchannel preferably comprising a reactant flow path that includesmultiple conduits) in the first reactor bed (7). In one embodiment, thechannels include one or more honeycomb monolith type structures.Honeycomb monoliths include extruded porous structures as are generallyknown in the reaction industry, such as in catalytic converters, etc.The term “honeycomb” is used broadly herein to refer to a porouscross-sectional shape that includes multiple flow paths or conduitsthrough the extruded monolith and is not intended to limit the structureor shape to any particular shape. The honeycomb monolith enables lowpressure loss transference while providing contact time and heattransfer. A mixer is preferably used between the zones (e.g., between orwithin a medium between the first and second reactors) to enable orassist combustion within and/or subsequent to the mixer. Each of thefirst channel and the second channel is defined broadly to mean therespective conductive conduit(s) or flow path(s) by which one of thereactants and synthesis gas flows through the first reactor bed (7) andmay include a single conduit or more preferably and more likely,multiple conduits (e.g., tens, hundreds, or even thousands ofsubstantially parallel conduits or tubes) that receive feed, such asfrom a gas/vapor distributor nozzle or dedicated reactant port.

The conduits each may have generally any cross-sectional shape, althougha generally circular or regular polygon cross-sectional shape may bepreferred. Each channel may preferably provide substantially parallel,generally common flow through the reactor media. Thus, a first channelmay be merely a single conduit, but more likely will be many conduits,(depending upon reactor size, flow rate, conduit size, etc.), forexample, such as exemplified in FIG. 2. A channel preferably includesmultiple conduits that each receive and conduct a reactant, such asdelivered by a nozzle in a gas distributor. The conduits may be isolatedfrom each other in terms of cross flow along the flow path (e.g. not influid communication), or they may be substantially isolated, such thatreactant permeation through a conduit wall into the adjacent conduit issubstantially inconsequential with respect to reactant flow separation.One preferred reactor embodiment includes multiple segments, wherebyeach segment includes a first channel and a second channel, such thatafter exiting the reactor, the respective first reactant is mixed withthe respective second reactant in a related mixer segment. Multiplesegments are included to provide good heat distribution across the fullcross-sectional area of the reactor system.

Referring to FIG. 4, mixer segment (45), for example, may mix thereactant flows from multiple honeycomb monoliths arranged within aparticular segment. Each monolith preferably comprises a plurality (morethan one) of conduits. The collective group of conduits that transmitthe first reactant may be considered the first channel and a particularreactor segment may include multiple collective groups of monolithsand/or conduits conducting the first reactant, whereby the segmentcomprising a channel for the first reactant. Likewise, the secondreactant may also flow through one or more monoliths within a segment,collectively constituting a second channel. Thus, the term “channel” isused broadly to include the conduit(s) or collective group of conduitsthat conveys at least a first or second reactant. A reactor segment mayinclude only a first and second channel or multiple channels formultiple flow paths for each of the first and second reactants. A mixersegment (45) may then collect the reactant gas from both or multiplechannels. Preferably, a mixer segment (45) will mix the effluent fromone first channel and one second channel.

It is recognized that in some preferred embodiments, some or evenseveral of the conduits within a channel will likely convey a mixture offirst and second reactants, due at least in part to some mixing at thefirst end (17) of the first reactor. However, the numbers of conduitsconveying combustible mixtures of first and second reactants issufficiently low such that the majority of the stoichiometricallyreactable reactants will not react until after exiting the second end ofthe first reactor. The axial location of initiation of combustion orexothermic reaction within those conduits conveying a mixture ofreactants is controlled by a combination of temperature, time, and fluiddynamics. Fuel and oxygen usually require a temperature-dependent andmixture-dependent autoignition time to combust. Still though, somereaction will likely occur within an axial portion of the conduitsconveying a mixture of reactants. However, this reaction is acceptablebecause the number of conduits having such reaction is sufficientlysmall that there is only an acceptable or inconsequential level ofeffect upon the overall heat balance within the reactor. The designdetails of a particular reactor system should be designed so as to avoidmixing of reactants within the conduits as much as reasonably possible.

The process according to the present invention requires no largepressure swings to cycle the reactants and products through the reactorsystem. In some preferred embodiments, the reforming or pyrolysis ofvolatized hydrocarbon step occurs at relatively low pressure, such asless than about 50 psia, while the regeneration step may also beperformed at similar pressures, e.g., less than about 50 psia, or atslightly higher, but still relatively low pressures, such as less thanabout 250 psia. In some preferred embodiments, the volatized hydrocarbonpyrolysis step is performed at a pressure of from about 5 psia to about45 psia, preferably from about 15 psia to about 35 psia. Ranges fromabout 7 psia to about 35 psia and from about 15 psia to about 45 psiaare also contemplated. The most economical range may be determinedwithout more than routine experimentation by one of ordinary skill inthe art in possession of the present disclosure. Pressures higher orlower than that disclosed above may be used, although they may be lessefficient. By way of example, if combustion air is obtained fromextraction from a gas turbine, it may be preferable for regeneration tobe carried out at a pressure of, for example, from about 100 psia toabout 250 psia. However, if by way of further example, the process ismore economical with air obtained via fans or blowers, the regenerationmay be carried out at lower pressures such as 15-45 psia. In oneembodiment of the present invention, the pressure of the pyrolysis andregeneration steps are essentially the same, the difference between thepressures of the two steps being less than about 10 psia.

It is understood that some method of flow control (e.g. valves, rotatingreactor beds, check valves, louvers, flow restrictors, timing systems,etc.) is used to control gas flow, actuation, timing, and to alternatephysical beds between the two flow systems. In the regeneration step,air and fuel must be moved through the reactor system and combined forcombustion. Air can be moved such as via compressor, blower, or fan,depending on the operating conditions and position desired for thereactor. If higher pressure air is used, it may be desirable to expandthe flue gas through an expansion turbine to recover mechanical energy.In addition, some fraction of exhaust gas may be recycled and mixed withthe incoming air. An exhaust gas recycle (EGR) stream may be suppliedwith at least one of the supplied first reactant and second reactant inthe first reactor. This EGR may be used to reduce the oxygen content ofthe regeneration feed, which can reduce the maximum adiabatic flametemperature of the regeneration feed. In the absence of EGR, CH4/airmixtures have a maximum adiabatic flame temperature of about 1980° C.;H₂/air mixtures are about 2175° C. Thus, even if average temperature iscontrolled by limiting the flow rate of fuel, any poor diluting couldresult in local hot spots that approach the maximum flame temperature.Use of EGR can reduce the maximum hot spot temperature by effectivelyincreasing the amount of diluent such as N₂ (and combustion products)that accompany the oxygen molecules.

For example, when 50 percent excess air is used for combustion, themaximum adiabatic flame temperature for H₂-fuel/air combustion decreasesfrom about 3947° F. (2175° C.) to about 2984° F. (1640° C.). Reducingthe oxygen content of the supplied air to about 13 percent would makeabout 2984° F. (1640° C.) the maximum adiabatic flame temperature,regardless of local mixing effects. The reforming or pyrolysis step andflow scheme is illustrated in FIG. 1( a). The vapor phase of theseparation of the hydrocarbon feed stream is transferred to the reactorsystem inlet, preferably mixed with or supplied with hydrogen or asource for hydrogen as a diluent or stripping agent, either within thesecond reactor or prior to entry into the second reactor, and ispyrolyzed in the high temperature heat bubble created by theregeneration step. After leaving the second reactor and the optionalmixer, the pyrolyzed product stream must be cooled or quenched to haltthe conversion process at the acetylene or other appropriate stage. Thetiming for this step is important because is the reaction is not timelyand properly quenched, some desired products, such as acetylene, will bepassed by the reaction and the pyrolysis product will not have thedesired selectivity to the valuable or desired products. Some pyrolysisproducts, however, are still rarely a desired final material for processexport. Rather, a preferred use for the produced pyrolysis products,such as acetylene, is as an intermediate product in a flow processwithin a chemical plant, in route to other preferred products, such asvinyl esters, ethylene, acetaldehyde, propanal, and/or propanol, acrylicacid, and so on. Typical desired pyrolysis products may be an olefinand/or an alkyne. Some commonly desired olefins may include ethylene,propylene, and/or butylene. Some commonly desired alkynes may includeacetylene.

After quenching, the synthesized gas stream may be provided to aseparation process that separates the acetylene, methane, hydrogen, andother gases. Recovered methane and hydrogen may be recycled forprocessing again in the reactor system. Separate process sequences mayconvert the acetylene to other final products. Each of these productsmay be further processed to provide yet additional useful products,e.g., acetaldehyde is typically an intermediate in the manufacture ofethanol, acetic acid, butanals, and/or butanols. Ethylene is a basicbuilding block of a plethora of plastics, and may typically be thepreferred use for the created acetylene, from the perspective of volumeand value. Ethylene is conveniently manufactured from acetylene byhydrogenation. In some embodiments of the invention, it may also be acoproduct of the inventive volatized hydrocarbon conversion process.Another product of high interest is ethanol, which may be convenientlymanufactured by first hydrating the acetylene to acetaldehyde and thenhydrogenating acetaldehyde to ethanol. Ethanol is of interest because itis easily transported from a remote location and is easily dehydrated toethylene. Ethanol may also be suitable for use as a motor fuel, if themanufacturing can be sufficiently low in cost.

Conversion of a volatized hydrocarbon stream to acetylene leaves asurplus of hydrogen. An idealized reaction is to crack the aliphaticchains into various methyl groups and continue the pyrolysis reactionvia further conversion of the methyls to acetylene. An exemplaryreaction for conversion of methane is:2CH₄→C₂H₂+3H₂ consuming about+45 kcal/mole of converted CH₄

As suggested by the above reaction, hydrogen is a valuable by-product ofthe present process. To a lesser extent, ethylene and propylene are alsovaluable products, produced as a result of incomplete reduction ofvolatized hydrocarbon to higher hydrocarbon. Unreacted volatizedhydrocarbon is also a valuable product for recovery. Accordingly,separation and recovery of hydrogen, separation and recovery of olefinssuch as ethylene and propylene, and separation and recovery ofunconverted volatized hydrocarbon vapor feed are each individually andcollectively preferred steps in the process according to the invention.Unconverted volatized hydrocarbon is preferably returned to thehydropyrolysis reactor so that it may be converted on a second pass. Anamount of hydrogen should also be returned to the hydropyrolysis reactorthat is sufficient to control the selectivity of the productdistribution.

Since hydrogen is created (not consumed) in the reforming pyrolysisreaction, it will be necessary to purge hydrogen from the process. Forexample, conversion of methane to acetylene, with subsequenthydrogenation to ethylene, will generate about one H₂ for every CH₄converted. Hydrogen has a heat of combustion of about 57 Kcal/mole H₂,so the hydrogen purged from the process has a heating value that is inthe range of what is needed as regeneration fuel. Of course, if there isan alternate, high-value use for the leftover hydrogen, then natural gascould be used for all or part of the regeneration fuel. But the leftoverhydrogen is likely to be available at low pressure and may possiblycontain methane or other diluents. Thus, use of hydrogen as regenerationfuel may also be an ideal disposition in a remote location. However,heavier feeds may not make as much excess hydrogen, and in someinstances may not make any appreciable volumes of excess hydrogen. Theamount of excess hydrogen generated will depend strongly on the overallhydrogen to carbon ratio of the feedstock and upon the desired ultimateproduct. For example, an ethylene product will result in less excesshydrogen than an ethyne (acetylene) product.

FIG. 2 illustrates another exemplary reactor system that may be suitablein some applications for controlling and deferring the combustion offuel and oxidant to achieve efficient regeneration heat. FIG. 2 depictsa single reactor system, operating in the regeneration cycle. Theinventive reactor system comprises two reactors zones or two reactors.The recuperator (27) is the zone primarily where quenching takes placeand provides substantially isolated flow paths or channels fortransferring both of the quenching reaction gases through the reactormedia, without incurring combustion until the gases arrive proximate orwithin the reactor core (13) in FIG. 1. The reformer (2) is the reactorwhere regeneration heating and volatized hydrocarbon reformationprimarily occurs, and may be considered as the second reactor forpurposes herein. Although the first and second reactors in the reactorsystem are identified as separately distinguishable reactors, it isunderstood and within the scope of the present invention that the firstand second reactors may be manufactured, provided, or otherwise combinedinto a common single reactor bed, whereby the reactor system might bedescribed as comprising merely a single reactor that integrates bothcycles within the reactor. The terms “first reactor” and “secondreactor” merely refer to the respective zones within the reactor systemwhereby each of the regeneration, reformation, quenching, etc., stepstake place and do not require that separate components be utilized forthe two reactors. However, most preferred embodiments will comprise areactor system whereby the recuperator reactor includes conduits andchannels as described herein, and the reformer reactor may similarlypossess conduits. Other preferred embodiments may include a reformerreactor bed that is arranged different from and may even includedifferent materials from, the recuperator reactor bed. The beddingarrangement of the reformer or second reactor may be provided as desiredor as prescribed by the application and no particular design is requiredherein of the reformer reactor, as to the performance of the inventivereactor system. Routine experimentation and knowledge of the volatizedhydrocarbon pyrolysis art may be used to determine an effectivereformer/second reactor design.

As discussed previously, the first reactor or recuperator (27) includesvarious gas conduits (28) for separately channeling two or more gasesfollowing entry into a first end (29) of the recuperator (27) andthrough the regenerative bed(s) disposed therein. A first gas (30)enters a first end of a plurality of flow conduits (28). In addition toproviding a flow channel, the conduits (28) also comprise effective flowbarriers (e.g., which effectively function such as conduit walls) toprevent cross flow or mixing between the first and second reactants andmaintain a majority of the reactants effectively separated from eachother until mixing is permitted. In a preferred embodiment of thepresent invention, the recuperator is comprised of one or more extrudedhoneycomb monoliths. A small reactor may include a single monolith,while a larger reactor can include a number of monoliths, while stilllarger reactor may be substantially filled with an arrangement of manyhoneycomb monoliths.

Honeycomb monoliths preferred in the present invention (which areadjacent a first end (9) of the first reactor (7)) can be characterizedas having open frontal area (or geometric void volume) between about 40%and 80%, and having conduit density between about 50 and 2000 pores persquare inch, more preferably between about 100 and 1000 pores per squareinch. (For example, in one embodiment, the conduits may have a diameterof only a few millimeters, and preferably on the order of about onemillimeter.) Reactor media components, such as the monoliths oralternative bed media, preferably provide for at least one of the firstand second channels and preferably both channels to include a packingwith an average wetted surface area per unit volume that ranges fromabout 50 ft⁻¹ to about 3000 ft⁻¹, more preferably from about 100 ft⁻¹ to2500 ft⁻¹, and still more preferably from about 200 ft⁻¹ to 2000 ft⁻¹,based upon the volume of the first reactor that is used to convey areactant. These wetted area values apply to the channels for both of thefirst and second reactants. These relatively high surface area per unitvolume values are likely preferred for many embodiments to aid achievinga relatively quick change in the temperature through the reactor, suchas generally illustrated by the relatively steep slopes in the exemplarytemperature gradient profile graphs, such as in FIGS. 1( a), 1(b), and6. The quick temperature change is preferred to permit relatively quickand consistent quenching of the reaction to prevent the reaction fromcontinuing and creating coke.

Preferred reactor media components also provide for at least one of thefirst and second channels in the first reactor and more preferably forboth channels, to include a packing that includes a high volumetric heattransfer coefficient (e.g., greater than or equal to 0.02 cal/cm³s° C.,preferably greater than about 0.05 cal/cm³s° C., and most preferablygreater than 0.10 cal/cm³s° C.), have low resistance to flow (lowpressure drop), have operating temperature range consistent with thehighest temperatures encountered during regeneration, have highresistance to thermal shock, and have high bulk heat capacity (e.g., atleast about 0.10 cal/cm³° C., and preferably greater than about 0.20cal/cm³° C.). As with the high surface area values, these relativelyhigh volumetric heat transfer coefficient value and other properties arealso likely preferred for many embodiments to aid in achieving arelatively quick change in the temperature through the reactor, such asgenerally illustrated by the relatively steep slopes in the exemplarytemperature gradient profile graphs, such as in FIGS. 1( a), 1(b), and6. The quick temperature change is preferred to permit relatively quickand consistent quenching of the reaction to prevent the reaction fromcontinuing too long and creating coke or carbon buildup. The citedvalues are averages based upon the volume of reactor used for conveyanceof a reactant.

Alternative embodiments may use reactor media other than the describedand preferred honeycomb monoliths, such as whereby the channelconduits/flow paths may include a more tortuous pathways (e.g.convoluted, complex, winding and/or twisted but not linear or tubular),than the previously described monoliths, including but not limited tolabyrinthine, variegated flow paths, conduits, tubes, slots, and/or apore structure having channels through a portion(s) of the reactor andmay include barrier portion, such as along an outer surface of a segmentor within sub-segments, having substantially no effective permeabilityto gases, and/or other means suitable for preventing cross flow betweenthe reactant gases and maintaining the first and second reactant gasessubstantially separated from each other while axially transiting therecuperator (27). For such embodiments, the complex flow path may createa lengthened effective flow path, increased surface area, and improvedheat transfer. Such design may be preferred for reactor embodimentshaving a relatively short axial length through the reactor. Axiallylonger reactor lengths may experience increased pressure drops throughthe reactor. However for such embodiments, the porous and/or permeablemedia may include, for example, at least one of a packed bed, anarrangement of tiles, a permeable solid media, a substantiallyhoneycomb-type structure, a fibrous arrangement, and a mesh-type latticestructure. It may be preferred that the media matrix provides highsurface area to facilitate good heat exchange with the reactant andproduced gases.

It may be preferred to utilize some type of equipment or method todirect a flow stream of one of the reactants into a selected portion ofthe conduits. In the exemplary embodiment of FIG. 2, a gas distributor(31) directs a second gas stream (32) to second gas stream channels thatare substantially isolated from or not in fluid communication with thefirst gas channels, here illustrated as channels (33). The result isthat at least a portion of gas stream (33) is kept separate from gasstream (30) during axial transit of the recuperator (27). In a preferredembodiment, the regenerative bed(s) of the recuperator zone comprisechannels having a gas or fluid barrier that isolates the first reactantchannels from the second reactant channels. Thereby, both of the atleast two reactant gases that transit the channel means may fullytransit the regenerative bed(s), to quench the regenerative bed, absorbheat into the reactant gases, before combining to react with each otherin the combustion zone.

By keeping the reactants (30) and (32) substantially separated, thepresent invention defers or controls the location of the combustion orother heat release that occurs due to exothermic reaction.“Substantially separated” means that at least 50 percent, preferably atleast 75 percent, and more preferably at least 90 percent of thereactant having the smallest or limiting stoichiometrically reactableamount of reactant, as between the first and second reactant streams,has not become consumed by reaction by the point at which these gaseshave completed their axial transit of the recuperator (27). In thismanner, the majority of the first reactant (30) is kept isolated fromthe majority of the second reactant (32), and the majority of the heatrelease from the reaction of combining reactants (30) and (32) will nottake place until the reactants begin exiting the recuperator (27).Preferably the reactants are gases, but some reactants may comprise aliquid, mixture, or vapor phase.

The percent reaction for these regeneration streams is meant the percentof reaction that is possible based on the stoichiometry of the overallfeed. For example, if gas (30) comprised 100 volumes of air (80 volumesN₂ and 20 Volumes O₂), and gas (32) comprised 10 volumes of Hydrogen,then the maximum stoichiometric reaction would be the combustion of 10volumes of hydrogen (H₂) with 5 volumes of Oxygen (O₂) to make 10volumes of H₂O. In this case, if 10 volumes of hydrogen were actuallycombusted in the recuperator zone (27), this would represent 100%reaction of the regeneration stream. This is despite the presence ofresidual un-reacted oxygen, because that un-reacted oxygen was presentin amounts above the stoichiometric requirement. Thus, the hydrogen isthe stoichiometrically limiting component. Using this definition, it ispreferred than less than 50% reaction, more preferred than less than 25%reaction, and most preferred that less than 10% reaction of theregeneration streams occur during the axial transit of the recuperator(27).

In a preferred embodiment, the channels (28) and (33) comprise materialsthat provide adequate heat transfer capacity to create the temperatureprofiles (4) and (8) illustrated in FIG. 1 at the space velocityconditions of operation. Adequate heat transfer rate is characterized bya heat transfer parameter ΔT_(HT), below about 500° C., more preferablybelow about 100° C. and most preferably below about 50° C. The parameterΔT_(HT), as used herein, is the ratio of the bed-average volumetric heattransfer rate that is needed for recuperation, to the volumetric heattransfer coefficient of the bed, h_(v). The volumetric heat transferrate (e.g. cal/cm³ sec) that is sufficient for recuperation iscalculated as the product of the gas flow rate (e.g. gm/sec) with thegas heat capacity (e.g. ca./gm° C.) and desired end-to-end temperaturechange (excluding any reaction, e.g. ° C.), and then this quantitydivided by the volume (e.g. cm³) of the recuperator zone (27) traversedby the gas. The ΔT_(HT) in channel (28) is computed using gas (30),channel (33) with gas (32), and total recuperator zone (27) with totalgas. The volumetric heat transfer coefficient of the bed, h_(v), istypically calculated as the product of a area-based coefficient (e.g.cal/cm²s° C.) and a specific surface area for heat transfer (av, e.g.cm²/cm³), often referred to as the wetted area of the packing.

In a preferred embodiment, channels (28) and (33) comprise ceramic(including but not limited to zirconia), alumina, or other refractorymaterial capable of withstanding temperatures exceeding 1200° C., morepreferably 1500° C., and still more preferably 1700° C. Materials havinga working temperature of up to and in excess of 2000° C. might bepreferred where there is concern with reaching the bed reactionadiabatic maximum temperature for sustained periods of time, to preventreactor bed damage, provided the project economics and conditionsotherwise permit use of such materials. In a preferred embodiment,channels (28) and (33) have wetted area between 50 ft⁻¹ and 3000 ft⁻¹,more preferably between 100 ft⁻¹ and 2500 ft⁻¹, and most preferablybetween 200 ft⁻¹ and 2000 ft⁻¹. Most preferably, channel means (28)comprise a ceramic honeycomb, having channels running the axial lengthof the recuperator reactor (27).

Referring again briefly to FIGS. 1( a) and 1(b), the inventive reactorsystem includes a first reactor (7) containing a first end (9) and asecond end (11), and a second reactor (1) containing a primary end (3)and a secondary end (5). The embodiments illustrated in FIGS. 1( a),1(b), and 2 are merely simple illustrations provided for explanatorypurposes only and are not intended to represent a comprehensiveembodiment. Reference made to an “end” of a reactor merely refers to adistal portion of the reactor with respect to an axial mid-point of thereactor. Thus, to say that a gas enters or exits an “end” of thereactor, such as end (9), means merely that the gas may enter or exitsubstantially at any of the various points along an axis between therespective end face of the reactor and a mid-point of the reactor, butmore preferably closer to the end face than to the mid-point.

With regard to the various exemplified embodiments, FIG. 3 illustratesan axial view of an exemplary gas distributor (31) having apertures(36). Referring to both FIGS. 2 and 3, apertures (36) may direct thesecond reactant gas (32) preferentially to select channels (33). In apreferred embodiment, apertures (36) are aligned with, but are notsealed to, the openings/apertures of select channels (33). Nozzles orinjectors (not shown) may be added to the apertures (36) that aresuitably designed to direct the flow of the second gas (32)preferentially into the select channels (33). By not “sealing” the gasdistributor apertures (36) (or nozzles/injectors) to the select channels(33), these channels may be utilized during the reverse flow or reactioncycle, increasing the overall efficiency of the system. Such “open” gasdistributor (31) may be preferred for many applications, over a “closed”system, to facilitate adaptation to multiple reactor systems, such aswhere the reactor/recuperator beds may rotate or otherwise move inrelation to the location of the gas stream for processing, e.g., such aswith a rotating bed type reactor system.

When a gas distributor nozzle or aperture (36) in an “open” systemdirects a stream of reactant gas (32) toward the associated inletchannel and associated conduits in the reactor (preferably a honeycombmonolith(s)), the contents of that stream of reactant gas (32) willtypically occupy a large number of honeycomb conduits (33) as ittraverses the recuperator. This outcome is a geometric result of thesize of the reactor segments and/or aperture size, relative to the sizeof the monolith honeycomb conduits. The honeycomb conduits occupied bygas (32) may, according to a preferred embodiment, be characterized as abundle of conduits, typically oriented along the same axis as theaperture (36) and its issuing stream of gas (32). Conduits located nearthe center of this bundle/channel will contain a high purity of gas (32)and thus will likely not support exothermic reaction. Conduits locatednear the outer edge of the bundle will be in close proximity to conduits(28) carrying the other reactant. In an “open” system as describedabove, some mixing of the first gas (30) and the second gas (32) will beunavoidable near the peripheral edges of each stream of gas (32) thatissues from the apertures (36). Thus, some conduits (28) and (33) nearthe outer edge of the bundle will carry some amount of both the firstgas (30) and the second gas (32). Reaction or combustion between gases(30) and (32) could happen in these conduits before the gases completelytraverse recuperator (27). Such gases would still be considered to besubstantially separated, as long as the resulting reaction of theregeneration streams within the recuperator (27) is less than 50%,preferably than less than 25%, and most preferably less than 10% of thestoichiometrically reactive reactant having the smallest or reactionlimiting presence.

In some alternative embodiments, the recuperator reactor (27) mayinclude, for example, packed bed or foam monolith materials (not shown)that permit more mixing or dispersion of reactants before fullytraversing the first reactor. In this case, additional reaction mayoccur in the recuperator (27) due to mixing within the recuperator thatis due to the axial dispersion of gases (30) and (32) as they passthough. This may still be an acceptable arrangement as long as themixing and subsequent reaction of the regeneration streams within therecuperator (27) is less than 50%, preferably than less than 25%, andmost preferably less than 10%. Methods for calculation of radialdispersion and mixing in bed media is known in the art.

During regeneration, the first gas (30) and second gas (32) transit therecuperator zone (27) via channels (28) and (33). It is a key aspect ofthis invention that heat, stored in the recuperator zone from theprevious quench cycle, is transferred to both the first and second gasesduring the regeneration cycle. The heated gases are then introduced intomixer (44). The gas mixer (44), located between the recuperator (27) andthe reactor (21), functions to mix the regenerating reaction gas streams(30) and (32), preferably at or near the interface of the reaction zone(21) and the mixer (44).

The mixer (44) is preferably constructed or fabricated of a materialable to withstand the high temperatures expected to be experienced inthe reaction zone during volatized hydrocarbon reforming at highselectivity and high conversion rates (>50 weight percent). In apreferred embodiment, mixer (44) is constructed from a material able towithstand temperatures exceeding 2190° F. (1200° C.), more preferably2730° F. (1500° C.), and most preferably 3090° F. (1700° C.). In apreferred embodiment, mixer means (34) is constructed of ceramicmaterial(s) such as alumina or silicon carbide for example.

FIG. 4 illustrates an axial view of one configuration of the mixer (44),together with a cut-away view FIG. 4 a, of one exemplary embodiment ofswirl-type mixer (47). The exemplary mixer (44) comprises mixer segments(45) having swirl mixer (47) located within the sections (45). In apreferred embodiment, mixer segments (45) are substantially equal incross sectional area and the swirl mixers (47) are generally centrallylocated within the sections (45). Mixer segments (45) are positionedwith respect to the reactor system to segment the gas flow of aplurality of gas channels (28) and (33). In a preferred embodiment,segments (45) may each have substantially equal cross sectional area tofacilitate intercepting gas flow from a substantially equal number ofgas channel means (28) and (33). Also in a preferred embodiment, the gaschannels (28) and (33) are distributed within recuperator reactor (27)such that each of the segments (45) intercepts gas flow from asubstantially equal fraction of both first gas channel means (28) andsecond gas channel means (33). Expressed mathematically, one can definef_(Ai) as the fraction of total cross sectional area encompassed bysection i, f_(28i) as the fraction of total channel means (28)intercepted by section i, and f_(33i) as the fraction of total channelmeans (33) intercepted by section i. In a preferred embodiment, for eachsection i, the values f_(28i), and f_(33i) will be within about 20percent of (i.e. between about 0.8 and 1.2 times) the value of f_(Ai),and more preferably within about 10%. One can further define f_(30i) asthe fraction of gas stream (30) intercepted by section i, and f_(32i) asthe fraction of gas stream (32) intercepted by the section i. In a morepreferred embodiment, for each section i, the values of f_(30i), andf_(32i) will be within about 20 percent of f_(Ai), and more preferablywithin about 10%.

FIG. 4 a illustrates an exemplary cut out section of an individual gasmixer segment (45) with swirl mixer (47). While the present inventionmay utilize a gas mixer known to the skilled artisan to combine gasesfrom the plurality of gas channel means (28) and (33), a preferredembodiment of this invention minimizes open volume of the gas mixer (44)while maintaining sufficient mixing and distribution of the mixed gases.The term open volume means the total volume of the swirl mixers (47) andgas mixer segment (45), less the volume of the material structure of thegas mixer. Accordingly, gas mixer segment (45) and swirl mixer (47) arepreferably configured to minimize open volume while concurrentlyfunctioning to provide substantial gas mixing of the gases exiting gaschannels (28) and (33). In a preferred practice of the invention, gasmixer segment (45) dimensions L and D, are tailored to achievesufficient mixing and distribution of gases (31) and (32) whileminimizing open volume. Dimension ratio L/D is preferably in the rangeof 0.1 to 5.0, and more preferably in the range of 0.3 to 2.5. Forgeneral segments of area A, a characteristic diameter D can be computedas 2(A/π)^(1/2).

In addition, the total volume attributable to the gas mixer (44) ispreferably tailored relative to the total volume of the first reactorbed (27) and reforming bed (21). Gas mixer (44) preferably has a totalvolume of less than about 20%, and more preferably less than 10% of thecombined volume of the recuperator zone (27), the reformation zone (21),and the gas mixer (44).

Referring again to FIG. 2, the mixer (44) as configured combines gasesfrom channels (33) and (28), and redistributes the combined gas acrossand into reaction zone (21). In a preferred embodiment, first reactantand second reactant are each a gas and one comprises a fuel and theother an oxidant. Fuel may comprise hydrogen, carbon monoxide,hydrocarbons, oxygenates, petrochemical streams, or mixtures thereof.Oxidant typically comprises a gas containing oxygen, commonly mixed withnitrogen, such as air. Upon mixing, the fuel and oxidant at mixer (44),the gases combust, with a substantial proportion of the combustionoccurring proximate to the entrance to the reaction zone (21).

The combustion of the fuel and oxygen-containing gas proximate to theentrance of the reformer or reaction zone (21) creates a hot flue gasthat heats (or re-heats) the reaction zone (21) as the flue gas travelsacross that zone. The composition of the oxygen-containing gas/fuelmixture is adjusted to provide the desired temperature of the reactionzone. The composition and hence reaction temperature may be controlledby adjusting the proportion of combustible to non-combustible componentsin the mixture. For example, non-combustible gases or other fluids suchas H₂O, CO₂, and N₂ also may be added to the reactant mixture to reducecombustion temperature. In one preferred embodiment, non-combustiblegases comprise steam, flue gas, or oxygen-depleted air as at least onecomponent of the mixture.

Referring again to regeneration FIG. 1( b), the reacted, hot combustionproduct passes through reformer (1), from the secondary end (5) to theprimary end (3), before being exhausted via conduit (18). The flow ofcombustion product establishes a temperature gradient, such asillustrated generally by example graph (8), within the reformation zone,which gradient moves axially through the reformation reaction zone. Atthe beginning of the regeneration step, this outlet temperature maypreferably have an initial value substantially equal (typically within25° C.) to the inlet temperature of the reforming feed of the preceding,reforming, step. As the regeneration step proceeds, this outlettemperature will increase somewhat as the temperature profile movestoward the outlet, and may end up 50° C. to 200° C. above the initialoutlet temperature. Preferably, the heated reaction product in theregeneration step heats at least a portion of the second reactor,preferably the secondary end (5) of the second reactor (1), to atemperature of at least about 1500° C., and more preferably to atemperature of at least about 1600° C. and still more preferably in someprocesses to a temperature of at least about 1700° C. Temperature andresidency time are both relevant to controlling the reaction speed andproduct.

Reactor system cycle time includes the time spent at regeneration plusthe time spent at reforming, plus the time required to switch betweenregeneration and reformation and vice versa. Thus, a half cycle may bethe substantially the time spent only on regeneration, or the time spendon reformation. A complete cycle includes heating the bed, feeding thevolatized hydrocarbon, and quenching the acetylene containing reactionproduct. Typical cycle times for preferred embodiments utilizinghoneycomb monoliths may be between 1 second and 240 seconds, althoughlonger times may be desired in some alternative embodiments. Morepreferably for the preferred monolith embodiments, cycle times may bebetween 2 seconds and 60 seconds. It is not necessary that theregeneration and reformation steps to have equal times, and in awell-refined application it is likely that these two times will not beequal.

As discussed above, in one preferred aspect, the inventive processincludes a process for pyrolyzing a hydrocarbon feedstock containingnonvolatiles in a regenerative pyrolysis reactor system, said processcomprising: (a) heating the nonvolatile-containing hydrocarbon feedstockupstream of a regenerative pyrolysis reactor system to a temperaturesufficient to form a vapor phase that is essentially free ofnonvolatiles and a liquid phase containing the nonvolatiles; (b)separating said vapor phase from said liquid phase; (c) feeding theseparated vapor phase and methane to the pyrolysis reactor system; and(d) converting the methane and separated vapor phase in said pyrolysisreactor system to form a pyrolysis product.

In another aspect, the invention includes a process for the manufactureof a hydrocarbon pyrolysis product from a hydrocarbon feed using aregenerative pyrolysis reactor system, wherein the reactor systemincludes (i) a first reactor comprising a first end and a second end,and (ii) a second reactor comprising primary end and a secondary end,and the first and second reactors are oriented in a series flowrelationship with respect to each other such that the secondary end ofthe second reactor is proximate the second end of the first reactor, theprocess comprises the steps of: (a) heating a nonvolatile-containinghydrocarbon feedstock upstream of the regenerative pyrolysis reactorsystem to a temperature sufficient to form a vapor phase that isessentially free of nonvolatiles and a liquid phase containing thenonvolatiles; (b) separating the vapor phase from the liquid phase; (c)supplying a first reactant through a first channel in the first reactorand supplying at least a second reactant through a second channel in thefirst reactor, such that the first and second reactants are supplied tothe first reactor from the first end of the first reactor; (d) combiningthe first and second reactants at the second end of the first reactorand reacting the combined reactants to produce a heated reactionproduct; (e) passing the heated reaction product through the secondreactor to transfer heat from the reaction product to the second reactorto produce a heated second reactor; (f) transferring at least a portionof the separated vapor phase from step (b) and methane to the pyrolysisreactor system, whereby the separated vapor phase and methane comminglewith each other within the reactor system; (g) feeding the separatedvapor phase and methane through the heated second reactor to convert theseparated vapor phase and methane into a pyrolysis product; (h)quenching the pyrolysis product; and (i) recovering the quenchedpyrolysis product from the reactor system. As previously discussed, instep (f), the commingling of the methane and separated vapor phase maybe initiated within the reactor system or at substantially any desiredpoint upstream of the reactor, such as in the transfer line, theseparator, or even upstream of the separator.

The present invention also includes an apparatus for pyrolyzing ahydrocarbon feedstock containing nonvolatiles in a regenerativepyrolysis reactor system, said apparatus comprising: (a) a heater toheat a nonvolatile-containing hydrocarbon feedstock to a temperaturesufficient to form a vapor phase that is essentially free ofnonvolatiles and a liquid phase containing the nonvolatiles; (b) aseparator to separate the vapor phase from the liquid phase; (c) amethane feed to provide a methane feedstock; and (d) a regenerativepyrolysis reactor system to receive the separated vapor phase and themethane feedstock, to heat and convert the separated vapor phase andmethane in said pyrolysis reactor system to form a pyrolysis product.

In yet another embodiment, the present invention includes an apparatusfor the manufacture of a hydrocarbon pyrolysis product from ahydrocarbon feed using a regenerative pyrolysis reactor system, whereinthe reactor system comprises: (a) a heater to heat anonvolatile-containing hydrocarbon feedstock to a temperature sufficientto form a vapor phase that is essentially free of nonvolatiles and aliquid phase containing the nonvolatiles; (b) a separator to separatethe vapor phase from the liquid phase; (c) a methane feed to provide amethane feedstock; and (d) a regenerative pyrolysis reactor system toreceive the separated vapor phase and the methane feedstock and convertthe separated vapor phase and methane in said pyrolysis reactor systemto form a pyrolysis product, the regenerative pyrolysis reactor systemincluding; (i) a first reactor comprising a first end and a second end;and (ii) a second reactor comprising primary end and a secondary end,the first and second reactors are oriented in a series flow relationshipwith respect to each other, wherein the first reactor comprises a firstchannel for conveying a first reactant through the first reactor and asecond channel for conveying a second reactant through the reactor.

While the present invention has been described and illustrated withrespect to certain embodiments, it is to be understood that theinvention is not limited to the particulars disclosed and extends to allequivalents within the scope of the claims.

1. A process for pyrolyzing a hydrocarbon feedstock containingnonvolatiles in a regenerative pyrolysis reactor system, said processcomprising: (a) heating the nonvolatile-containing hydrocarbon feedstockupstream of a regenerative pyrolysis reactor system to a temperaturesufficient to form a vapor phase and a nonvolatile-containing liquidphase; (b) separating said vapor phase from said liquid phase; (c)feeding the separated vapor phase and methane to the regenerativepyrolysis reactor system; wherein the molar ratio of methane toseparated vapor phase is from about 1:1 to about 5:1 and (d) convertingthe separated vapor phase in said regenerative pyrolysis reactor systemto form a pyrolysis product.
 2. The process according to claim 1,wherein the separated vapor phase is substantially free of nonvolatiles.3. The process according to claim 1, further comprising the step ofquenching the converted separated vapor phase to form the pyrolysisproduct.
 4. The process according to claim 1, wherein said hydrocarbonfeedstock is heated in step (a) to a temperature in the range of fromabout 200° C. to about 650° C.
 5. The process according to claim 1,wherein the pyrolysis reactor system comprises a reverse flowregenerative pyrolysis reactor system.
 6. The process according to claim1, wherein the nonvolatile-containing hydrocarbon feedstock comprises aliquid phase and at least 50 weight percent of the liquid phasehydrocarbon feedstock is converted in step (a) into said vapor phase. 7.The process according to claim 1, wherein said hydrocarbon feedstockcontaining nonvolatiles has a nominal end boiling point of at leastabout 200° C.
 8. The process according to claim 1, further comprisingthe step of feeding at least one of another hydrocarbon feed, hydrogen,and steam to at least one of (i) a unit that performs at least a portionof the separation in step (b), and (ii) the pyrolysis reactor system. 9.The process according to claim 1, wherein separation step (b) comprisesuse of at least one of a distillation column, flash drum, knock-outdrum, flash drum having heating means within the drum, knock-out drumhaving a heating means within the knock-out drum, and combinationsthereof.
 10. The process according to claim 1, further comprising thestep of maintaining a temperature of the vapor phase within theseparation vessel during the separation of said vapor phase from saidliquid phase, at a constant temperature between about 200° C. to about650° C.
 11. The process according to claim 1, wherein the regenerativepyrolysis reactor system heats the separated vapor phase and the methaneto a temperature of at least about 1200° C. to convert the separatedvapor phase to the pyrolysis product.
 12. The process according to claim1, wherein the regenerative pyrolysis reactor system heats the separatedvapor phase and the methane to a temperature of at least about 1400° C.to convert the separated vapor phase and methane to the pyrolysisproduct.
 13. A process for the manufacture of a hydrocarbon pyrolysisproduct from a hydrocarbon feed using a regenerative pyrolysis reactorsystem, performed using a regenerative pyrolysis reactor system, whereinthe reactor system includes (i) a first reactor comprising a first endand a second end, and (ii) a second reactor comprising primary end and asecondary end, and the first and second reactors are oriented in aseries flow relationship with respect to each other such that thesecondary end of the second reactor is proximate the second end of thefirst reactor, said process comprises the steps of: (a) heating anonvolatile-containing hydrocarbon feedstock upstream of theregenerative pyrolysis reactor system to a temperature sufficient toform a vapor phase that is essentially free of nonvolatiles and a liquidphase containing nonvolatiles; (b) separating the vapor phase from theliquid phase; (c) supplying a first reactant to a first end of the firstreactor and through a first channel in the first reactor and supplyingat least a second reactant to the first end of the first reactor andthrough a second channel in the first reactor; (d) combining the firstand second reactants at the second end of the first reactor and reactingthe combined first and second reactants to produce a heated reactionproduct; (e) passing the heated reaction product into and through thesecond reactor to transfer heat from the reaction product to the secondreactor to produce a heated second reactor; (f) feeding the separatedvapor phase and methane through the heated second reactor whereby theseparated vapor phase and methane are commingled with each other withinthe reactor system to thermally convert the separated vapor phase into apyrolysis product; wherein the molar ratio of methane to separated vaporphase is from about 1:1 to about 5:1 (g) quenching the pyrolysisproduct; and (h) recovering the quenched pyrolysis product from thereactor system.
 14. The process of claim 13, further comprising the stepof quenching the pyrolysis product in the first reactor to cool thepyrolysis product and heat the first reactor.
 15. The process of claim13, further comprising the step of preventing at least a majority of thefirst reactant from reacting with the second reactant within the firstreactor.
 16. The process of claim 13, further comprising the step offurther supplying a diluent in the second reactor, substantiallysimultaneously with step (f) of feeding the separated vapor phase andmethane through the heated second reactor.
 17. The process according toclaim 13, wherein said heating of said hydrocarbon feedstock in step (a)is carried out by at least one of a heat exchanger, steam injection, thereactor system, a fired heater, and combinations thereof.
 18. Theprocess according to claim 1, further comprising the step ofhydrogenating said pyrolysis product to form an olefin pyrolysisproduct.
 19. The process according to claim 1, wherein the regenerativepyrolysis reactor system heats the separated vapor phase and the methaneto a temperature of at least about 1700° C. to convert the separatedvapor phase and methane to the pyrolysis product.
 20. The processaccording to claim 13, wherein the first channel and the second channelin the first reactor are different channels in a reactor bed.